I. Field of the Invention
This invention relates to a process for reforming with sulfur-sensitive, polymetallic platinum-containing catalysts wherein sulfur is added to an intermediate or final reactor, or reaction zone, or zones, and excluded from the initial reactor, or reaction zone, of the series.
II. Background and Problems
Catalytic reforming, or hydroforming, is a well established industrial process employed by the petroleum industry for improving the octane quality of naphthas or straight run gasolines. In reforming, a multi-functional catalyst is employed which contains an acid component and a metal hydrogenation-dehydrogenation (hydrogen transfer) component, or components, substantially atomically dispersed upon the surface of a porous, inorganic oxide support, notably alumina. Noble metal catalysts, notably of the platinum type, are currently employed as metal hydrogenation-dehydrogenation components, reforming being defined as the total effect of the molecular changes, or hydrocarbon reactions, produced by dehydrogenation of cyclohexanes and dehydroisomerization of alkylcyclopentanes to yield aromatics; dehydrogenation of paraffins to yield olefins; dehydrocyclization of paraffins and olefins to yield aromatics; isomerization of n-paraffins; isomerization of alkylcycloparaffins to yield cyclohexanes; isomerization of substituted aromatics; and hydrocracking of paraffins which produces gas, and inevitably coke, the latter being progressively deposited on the catalyst as reforming is continued.
In a reforming operation, one or a series of reactors, or a series of reaction zones, are employed. Typically, a series of reactors are employed, e.g., three or four reactors, these constituting the heart of the reforming unit. Each reforming reactor is generally provided with a fixed bed, or beds, of the catalyst which receive downflow feed, and each is provided with a preheater or interstage heater. A naphtha feed, with hydrogen, or recycle hydrogen gas, is concurrently passed through a preheat furnace and reactor, and then in sequence through subsequent interstage heaters and reactors of the series. The product from the last reactor is separated into a liquid fraction, and a vaporous effluent. The former is recovered as a C.sub.5.sup.+ liquid product. The latter is a gas rich in hydrogen, and usually contains small amounts of normally gaseous hydrocarbons, from which hydrogen is separated and recycled as "recycle gas" to the process to minimize coke production. In conventional operations, the recycle gas, which generally contains moisture and hydrogen sulfide impurities, is passed through a recycle gas drier which removes much of the moisture and hydrogen sulfide prior to the introduction of the recycle gas into the first reactor of the series.
The sum-total of the reforming reactions, supra, occurs as a continuum between the first and last reactor of the series, i.e., as the feed enters and passes over the first fixed catalyst bed of the first reactor and exits from the last fixed catalyst bed of the last reactor of the series. The reactions which predominate between the several reactors differ dependent principally upon the nature of the feed, and the temperature employed within the individual reactors. In the initial reaction zone, or first reactor, which is maintained at a relatively low temperature, the primary reaction involves the dehydrogenation of naphthenes to produce aromatics. The isomerization of naphthenes, notably C.sub.5 and C.sub.6 naphthenes, also occurs to a considerable extent. Most of the other reforming reactions also occur, but only to a lesser, or smaller extent. There is relatively little hydrocracking, and very little olefin or paraffin dehydrocyclization occurs in the first reactor. Naphthene dehydrogenation is an endothermic reaction, and consequently the reactions in the first reactor are extremely endothermic, generally accounting for as much as 2/5 to 3/5 of the total observed temperature difference (.DELTA.T) across the several catalyst beds contained in the several reactors of the series. Within the intermediate reactor(s), or reaction zone(s), the temperature is maintained somewhat higher than in the first, or lead reactor of the series, and it is believed that the primary reactions in the intermediate reactor, or reactors, involve the isomerization of naphthenes and paraffins. Where, e.g., there are two reactors disposed between the first and last reactor of the series, it is believed that the principal reaction involves the isomerization of naphthenes, normal paraffins and isoparaffins. Some dehydrogenation of naphthenes may, and usually does occur, at least within the first of the intermediate reactors. There is usually some hydrocracking, at least more than in the lead reactor of the series, and there is more olefin and paraffin dehydrocyclization. The net effect of the reactions which occur in this reactor are endothermic, and though the temperature drop between the feed inlet and feed outlet is not as large as in that of the initial reactor (even though the second reactor generally contains a larger catalyst charge than the initial reactor), it is nonetheless considerable. The third reactor of the series, or second intermediate reactor, is generally operated at a somewhat higher temperature than the second reactor of the series. It is believed that the naphthene and paraffin isomerization reactions continue as the primary reaction in this reactor, but there is very little naphthene dehydrogeneration. There is a further increase in paraffin dehydrocyclization, and more hydrocracking. The net effect of the reactions which occur in this reactor is also endothermic, though the temperature drop between the feed inlet and feed outlet is smaller than in the first two reactors. In the final reaction zone, or final reactor, which is typically operated at the highest temperature of the series, it is believed that paraffin dehydrocyclization, particularly the dehydrocyclization of the short chain, notably C.sub.6 and C.sub.7 paraffins, is the primary reaction. The isomerization reactions continue, and there is more hydrocracking and coke formation in this reactor than in any of the other reactors of the series. The net effect of the reactions which occur in this reactor is also generally endothermic.
Platinum is widely commercially used in the production of reforming catalysts, and platinum-on-alumina catalysts have been commercially employed in refineries for the last few decades. In more recent years polymetallic catalysts have been used. These are catalysts wherein additional metallic components have been added to platinum as promoters to further improve the activity or selectivity, or both, of the basic platinum catalyst, e.g., iridium, rhenium, tin, and the like. Such catalysts possess superior activity, or selectivity, or both, as contrasted with the basic platinum catalyst. Platinum-rhenium catalysts by way of example possess admirable selectivity as contrasted with platinum catalysts, selectivity being defined as the ability of the catalyst to produce high yields of C.sub.5.sup.+ liquid products with concurrent low production of normally gaseous hydrocarbons, i.e., methane and other gaseous hydrocarbons, and coke.
The activity of the catalyst gradually declines due to the build-up of coke. Coke formation is believed to result from the deposition of coke precursors such as anthracene, coronene, ovalene, and other condensed ring aromatic molecules on the catalyst, these polymerizing to form coke. During operation, the temperature of the process is gradually raised to compensate for the activity loss caused by the coke deposition. Eventually, however, temperature increases cannot compensate for the loss in catalytic activity and hence it becomes necessary to reactivate the catalyst. Consequently, in all processes of this type the catalyst must necessarily be periodically regenerated by burning off the coke at controlled conditions.
Two major types of reforming are generally practiced in the multi reactor units, both of which necessitate periodic reactivation of the catalyst, the initial sequence of which requires regeneration, i.e., burning the coke from the catalyst. Reactivation of the catalyst is then completed in a sequence of steps wherein the agglomerated metal hydrogenation-dehydrogenation components are atomically redispersed. In the semi-regenerative process, a process of the first type, the entire unit is operated by gradually and progressively increasing the temperature to maintain the activity of the catalyst caused by the coke deposition, until finally the entire unit is shut down for regeneration, and reactivation, of the catalyst. In the second, or cyclic type of process, the reactors are individually isolated, or in effect swung out of line by various manifolding arrangements, motor operated valving and the like. The catalyst is regenerated to remove the coke deposits, and then reactivated while the other reactors of the series remain on stream. A "swing reactor" temporarily replaces a reactor which is removed from the series for regeneration and reactivation of the catalyst, until it is put back in series.
Change in the total, or overall, .DELTA.T is a good indication of changing performance in the reactors during an operating run, and correlates well with the ability of the reaction system to produce reformate octane value; which normally decreases throughout the run. In cyclic operations, in particular, the decline in the temperature drop across a catalyst bed is sometimes used as a criterion for selecting the next reactor candidate for regeneration of its catalyst charge.
Essentially all petroleum naphtha or synthetically derived naphtha feeds contain sulfur, a well known catalyst poison which can gradually accumulate upon and poison reforming catalysts. Most of the sulfur, because of this adverse effect, is removed from feed naphthas, e.g., by hydrofining and subsequent contact with guard beds packed or filled with sulfur adsorbents. The polymetallic reforming catalysts are particularly sulfur-sensitive, and particularlly susceptible to sulfur poisoning. The presence of even small and virtually infinitesimal amounts of sulfur in either the naphtha feed or hydrogen recycle gas, or both, it has been observed, adversely affects the process, and performance of the catalysts. Whereas various improvements have been made in adsorbents, and in the operation of guard beds to eliminate sulfur from the naphtha feed and hydrogen recycle gas, the complete elimination of sulfur from the naphtha does not appear practical, if indeed possible, and sulfur inevitably appears in the process. The effect of sulfur in the naphtha, even in concentration ranging only a few parts per million is that, in the overall reforming operation, the yield of hydrogen, aromatics, and C.sub.5.sup.+ liquid yield decreases as sulfur builds up and increases in the system, there is an increase in the rate of catalyst deactivation, and in the total production of C.sub.1 -C.sub.4 light gases.
III. Objects
It is, accordingly, the primary objective of this invention to provide a new and improved process useful in the operation of reforming units which employ highly active sulfur-sensitive polymetallic platinum-containing catalyst to produce high octane gasolines.
A specific object is to provide a novel process as characterized, but particularly a start-up procedure which will provide improved hydrogen purity and aromatics production, increased C.sub.5.sup.+ liquid yield, and decreased C.sub.4.sup.- light gas production in the operation of a reforming unit.
IV. The Invention
These objects and others are achieved in accordance with the present invention, embodying a process wherein, in a series of reforming zones, or reactors, each of which contains a bed, or beds of catalyst, the catalyst in each of which is constituted of a sulfur-sensitive, polymetallic platinum-containing catalyst which contains little or no coke, and naphtha, and hydrogen, are introduced into the lead reactor, passed in series from one reactor to another, and reacted at reforming conditions, sulfur is introduced into the final or tail reactor of the series to provide and maintain sulfur in concentration within the naphtha to said final reactor ranging from about 0.5 parts per million, based on the weight of the naphtha feed (wppm), to about 20 wppm, preferably from about 0.5 wppm to about 8 wppm, while excluding sulfur from the lead reactor of the series. The sulfur can be added to an intermediate reactor of the series, or added directly to the final reactor, preferably the latter. It has been found that the introduction, or addition of sulfur to the final reactor, or reaction zone of the series, during start-up, or that early portion of the operating cycle when the catalyst contains little or no coke, relative to the total operating cycle throughout which coke gradually, and progressively builds up and accumulates on the catalyst, provides improved hydrogen, aromatics, and C.sub.5.sup.+ liquid yields, and reduced C.sub.1 -C.sub.4 light gas make.
In accordance with this invention, sulfur is excluded, or its presence minimized, during start-up with a fresh or regenerated catalyst from the lead reactor of the series wherein naphthene dehydrogenation is the predominant reaction. A higher level of sulfur is maintained during this period in the final reactor of the series wherein paraffin and olefin dehydrocyclization are the predominant reactions. In carrying out such operation, sulfur is injected during this period into an intermediate or the tail reactor, preferably the latter, as hydrogen sulfide, or compound decomposable in situ to form hydrogen sulfide, in amount sufficient to provide sulfur in concentration ranging from about 0.5 wppm to about 20 wppm, preferably 0.5 wppm to about 8 wppm, based on naphtha feed. Sulfur is then removed from the product hydrogen gas from the last reactor of the series, suitably by passage of the gas through a sulfur adsorbent, and the gas is recycled. The gas is recycled to the lead reactor where it is added with fresh essentially sulfur-free naphtha. The essentially sulfur-free naphtha feed, and hydrogen, entering the lead reactor will provide a concentration of less than 0.5 wppm within the naphtha feed to the first reactor, and preferably no more than about 0.1 wppm sulfur. It is found that hydrogen, aromatics and C.sub.5.sup.+ liquid yields are improved, and reduced C.sub.1 -C.sub.4 light gas make up to such point in time that the catalyst in the final reactor, or reaction zone, contains no more than about 10 percent coke, based on the total weight of the catalyst, deposited thereon; and preferably no more than about 5 percent coke deposited upon the catalyst. This period corresponds generally from about 5 to about 60 percent of the operating cycle, and preferably from about 5 percent to about 30 percent of the total operating cycle which begins at start-up, or when the catalyst is first placed on-oil.
The catalyst employed in accordance with this invention is necessarily constituted of composite particles which contain, besides a carrier or support material, a hydrogenation-dehydrogenation component, or components, a halide component and, preferably, the catalyst is sulfided. The support material is constituted of a porous, refractory inorganic oxide, particularly alumina. The support can contain, e.g., one or more of alumina, bentonite, clay, diatomaceous earth, zeolite, silica, activated carbon, magnesia, zirconia, thoria, and the like; though the most preferred support is alumina to which, if desired, can be added a suitable amount of other refractory carrier materials such as silica, zirconia, magnesia, titania, etc., usually in a range of about 1 to 20 percent, based on the weight of the support. A preferred support for the practice of the present invention is one having a surface area of more than 50 m.sup.2 /g, preferably from about 100 to about 300 m.sup.2 /g, a bulk density of about 0.3 to 1.0 g/ml, preferably about 0.4 to 0.8 g/ml, an average pore volume of about 0.2 to 1.1 ml/g, preferably about 0.3 to 0.8 ml/g, and an average pore diameter of about 30 to 300 .ANG..
The metal hydrogenation-dehydrogenation component, or components, includes platinum, and one or more of iridium, rhenium, palladium, rhodium, tin, and tungsten. Preferably, the hydrogenation-dehydrogenation component, or components, are platinum and iridium or rhenium, or platinum and both iridium and rhenium. The hydrogenation-dehydrogenation component, or components, can be composited with or otherwise intimately associated with the porous inorganic oxide support or carrier by various techniques known to the art such as ion-exchange, coprecipitation with the alumina in the sol or gel form, and the like. For example, the catalyst composite can be formed by adding together suitable reagents such as a salt of platinum and a salt of rhenium and ammonium hydroxide or carbonate, and a salt of aluminum such as aluminum chloride or aluminum sulfate to form aluminum hydroxide. The aluminum hydroxide containing the salts of platinum and rhenium can then be heated, dried, formed into pellets or extruded, and then calcined in nitrogen or other nonagglomerating atmosphere. The metal hydrogenation components can also be added to the catalyst by impregnation, typically via an "incipient wetness" technique which requires a minimum of solution so that the total solution is absorbed, initially or after some evaporation.
It is preferred to deposit the platinum and iridium or rhenium metals, or both, and additional metals used as promoters, if any, on a previously pilled, pelleted, beaded, extruded, or sieved particulate support material by the impregnation method. Pursuant to the impregnation method, porous refractory inorganic oxides in dry or solvated state are contacted, either alone or admixed, or otherwise incorporated with a metal or metals-containing solution, or solutions, and thereby impregnated by either the "incipient wetness" technique, or a technique embodying absorption from a dilute or concentrated solution, or solutions, with subsequent filtration or evaporation to effect total uptake of the metallic components.
Platinum in absolute amount is usually supported on the carrier within the range of from about 0.05 to 3 percent, preferably from about 0.2 to 1 percent, based on the weight of the catalyst (dry basis). Rhenium, in absolute amount, is also usually supported on the carrier in concentration ranging from about 0.05 to about 3 percent, preferably from about 0.3 to about 1 percent, based on the weight of the catalyst (dry basis). Iridium, or metal other than platinum and rhenium, when employed, is also added in concentration ranging from about 0.05 to about 3 percent, preferably from about 0.2 to about 1 percent, based on the weight of the catalyst (dry basis). The absolute concentration of each, of course, is preselected to provide the desired ratio of rhenium:platinum for a respective reactor of the unit, as heretofore expressed.
In compositing the metals with the carrier, essentially any soluble compound can be used, but a soluble compound which can be easily subjected to thermal decomposition and reduction is preferred, for example, inorganic salts such as halide, nitrate, inorganic complex compounds, or organic salts such as the complex salt of acetylacetone, amine salt, and the like. When, e.g., platinum is deposited on the carrier, platinum chloride, platinum nitrate, chloroplatinic acid, ammonium chloroplatinate, potassium chloroplatinate, platinum polyamine, platinum acetylacetonate, and the like, are preferably used.
To enhance catalyst performance in reforming operations, it is also required to add a halogen component to the catalysts, fluorine and chlorine being preferred halogen components. The halogen is contained on the catalyst within the range of 0.1 to 3 percent, preferably within the range of about 0.3 to about 1.5 percent, based on the weight of the catalyst. When using chlorine as halogen component, it is added to the catalyst within the range of about 0.2 to 2 percent, preferably within the range of about 0.5 to 1.5 percent, based on the weight of the catalyst. The introduction of halogen into catalyst can be carried out by any method at any time. It can be added to the catalyst during catalyst preparation, for example, prior to, following or simultaneously with the incorporation of the metal hydrogenation-dehydrogenation component, or components. It can also be introduced by contacting a carrier material in a vapor phase or liquid phase with a halogen compound such as hydrogen fluoride, hydrogen chloride, carbon tetrachloride, or the like.
The catalyst is dried by heating at a temperature above about 80.degree. F., preferably between about 150.degree. F. and 300.degree. F., in the presence of nitrogen or oxygen, or both, in an air stream or under vacuum. The catalyst is calcined at a temperature between about 500.degree. F. to 1200.degree. F., preferably about 500.degree. F. to 1000.degree. F., either in the presence of oxygen in an air stream or in the presence of an inert gas such as nitrogen.
Sulfur is a highly preferred component of the catalysts, the sulfur content of the catalyst generally ranging to about 0.2 percent, preferably from about 0.03 percent to about 0.15 percent, based on the weight of the catalyst (dry basis). A fresh reforming catalyst is generally sulfided prior to its being placed on-oil, and since sulfur is lost during reforming, the catalyst is again sulfided during catalyst regeneration. The sulfur can be added to the catalyst by conventional methods, suitably by breakthrough sulfiding of a bed of the catalyst with a sulfur-containing gaseous stream, e.g., hydrogen sulfide in hydrogen, performed at temperatures ranging from about 350.degree. F. to about 1050.degree. F. and at pressures ranging from about 1 to about 40 atmospheres for the time necessary to achieve breakthrough, or the desired sulfur level.
The feed or charge stock can be a virgin naphtha cracked naphtha, a naphtha from a coal liquefaction process, a Fischer-Tropsch naphtha, or the like. Typical feeds are those hydrocarbons containing from about 5 to 12 carbon atoms, or more preferably from about 6 to about 9 carbon atoms. Naphthas, or petroleum fractions boiling within the range of from about 80.degree. F. to about 450.degree. F., and preferably from about 125.degree. F. to about 375.degree. F., contain hydrocarbons of carbon numbers within these ranges. Typical fractions thus usually contain from about 15 to about 80 vol. % paraffins, both normal and branched, which fall in the range of about C.sub.5 to C.sub.12, from about 10 to 80 vol. % of naphthenes falling within the range of from about C.sub.6 to C.sub.12, and from 5 through 20 vol. % of the desirable aromatics falling within the range of from about C.sub.6 to C.sub.12.
The reforming runs are initiated by adjusting the hydrogen and feed rates, and the temperature and pressure to operating conditions. The run is continued at optimum reforming conditions by adjustment of the major process variables, within the ranges described below:
______________________________________ Major Operating Typical Process Preferred Process Variables Conditions Conditions ______________________________________ Pressure, psig 50-750 100-500 Reactor Temp., .degree.F. 800-1200 850-1000 Recycle Gas Rate, SCF/B 1000-10,000 1500-5000 Feed Rate, W/Hr/W 0.5-10 1-5 ______________________________________